Fixed bed reactor

ABSTRACT

The invention provides a system for designing, operating, monitoring and/or diagnosing a chemical reaction, particularly a hydrotreating process, using a fixed bed catalytic reactor.

FIELD OF THE INVENTION

The invention relates to the design, operation, and/or diagnosis of afixed bed reactor and to the use of said reactor in a process. In anembodiment, the invention is directed to a process for the hydrogenationof aldehydes and ketones to make alcohols, and the design, operationand/or diagnosis of fixed-bed, gas-liquid, downflow catalytic reactorused for said process.

BACKGROUND OF THE INVENTION

The fixed bed catalytic reactor is a well-known, elegant device forcarrying out a chemical reaction utilizing a catalyst. There are myriadadvantages associated with this type of reactor, such as: the apparatusis typically simple to design, there are no moving parts to wear out,the catalyst stays in the reactor, it is easy to separate the reactionmixture from the catalyst, heat can be added or removed by, for example,the addition of cold gas or liquid quench, internal or external heatexchanger(s), wall heat transfer (e.g., in the case of small diametertubes like bench scale units or multi-tube-bundle reactors), or thereactor can be operated adiabatically.

There are numerous configurations of fixed bed catalytic reactors, themost common of which is probably cocurrent gas-liquid downflow,described, for instance, by R. Gupta, in “Cocurrent Gas-Liquid Downflowin Packed Beds”, Chapter 19, of the Handbook of Fluids in Motion (1983).Other configurations include cocurrent upflow and countercurrentoperations.

Whatever the specific configuration, theoretically the fixed bedcatalytic reactor is expected to provide, among other attributes,sufficient volume and residence time to provide the desired conversion,provide sufficient mass transfer rate of reactants and products throughthe gas-liquid interface and through the liquid film on the surface ofcatalyst particles, provide effective use of the entire catalystparticle and active sites throughout the cross section of particles inthe bed, provide uniform flow distribution over entire width and lengthof bed to utilize all of the catalyst, provide conditions where gas andliquid phases remain homogeneously mixed and do not separate, allow forconditions where all the catalyst is adequately wetted such that bothreactants are present and heat is transferred effectively from all zonesin the reactor, provide an effective method for controlling temperaturein a safe operating window or effective range to maximize reactionselectivity, product quality, catalyst life, and the like. See forinstance H. Hofmann: “Multiphase Catalytic Packed Bed Reactors”, Catal.Rev. Sci. Eng. 17 (1978) 71-117. However, it is still a long-sought goalto achieve all of the aforementioned attributes in a commercial reactor.

An example of the type of process that can be carried out in such areactor is hydrogenation. Heterogeneous catalytic hydrogenationprocesses of various kinds are widely practiced on a commercial scaleand are used for hydrogenation of a wide variety of organic feedstocks.

Specific examples include hydrogenation of aldehydes and ketones toalcohols, of unsaturated hydrocarbons to saturated hydrocarbons, ofacetylene-derived chemicals to saturated materials, of unsaturated fattyacids to saturated fatty acids, of esters of unsaturated fatty acids toesters of partially or fully hydrogenated fatty acids, of nitrites toprimary amines, of certain sugars to polyhydroxyalcohols. Other examplesinclude the hydrogenation of quinones (for example the hydrogenation of2-ethylanthraquinone as a step in the production of hydrogen peroxide),the production of cyclohexanol from cyclohexanone, the production ofiso-propanol from acetone, and the hydrogenation of unsaturatedhydrocarbons such as in the production of cyclohexane from benzene.

Typical catalysts for such hydrogenation reactions include Group VIIImetal catalysts such as cobalt, nickel, rhodium, palladium and platinum(using the traditional CAS version of the Periodic Table; see Chemicaland Engineering News, 63(5) 27, 1985), and also other metals such ascopper, zinc, and molybdenum.

Production of butane-1,4-diol by hydrogenation of but-2-yn-1,4-diol isan example of hydrogenation of an acetylene-derived chemical; a suitablecatalyst for this reaction has been described as a granularnickel-copper-manganese on silica gel. The production of stearic acid bycatalytic hydrogenation of the corresponding unsaturated acids, linoleicacid and linolenic acid, using a nickel, cobalt, platinum, palladium,chromium or copper/zinc catalyst, is an example of the hydrogenation ofunsaturated fatty acids to yield saturated fatty acids. So-called“hardening” of vegetable oils is an example of hydrogenation of estersof unsaturated fatty acids. Production of beta-phenylethylamine byhydrogenation of benzyl cyanide is an example of hydrogenation of anitrile. As examples of hydrogenation of sugars to polyhydroxyalcoholsthere can be mentioned hydrogenation of ketose and aldose sugars tohexahydroxyalcohols, for example hydrogenation of D-glucose to sorbitoland of D-mannose to mannitol.

An important route to C₃ and higher alcohols involves hydroformylationof olefins, such as ethylene, propylene, and butene-1, to yield thecorresponding aldehyde having one more carbon atom than the startingolefin, followed by hydrogenation to the alcohol. The commerciallyimportant Oxo Process comprises such a hydroformylation process,followed by hydrogenation. Thus, hydroformylation of ethylene yieldspropionaldehyde and propylene yields a mixture of n- andiso-butyraldehyde, followed by catalytic hydrogenation to thecorresponding alcohols, e.g. n-propanol and n-butanol. The importantplasticiser alcohol 2-ethylhexanol may be made, for instance, byalkali-catalyzed condensation of n-butyraldehyde to yield theunsaturated aldehyde, 2-ethyl-hex-2-enal, which is then hydrogenated toyield the desired 2-ethylhexanol. Historically the preferred catalystsfor such aldehyde hydrogenation reactions are the Group VIII metalcatalysts, particularly nickel, palladium, platinum, or rhodium.Numerous other systems have been proposed. The Oxo Process andvariations thereon are the subject of numerous patents and patentapplications, more recent examples of which are WO2003083788A2 andWO2003082789A2, and which in turn recite numerous references to the samesubject matter.

Hydrodesulphurisation is another commercially important hydrogenationreaction. This is the removal of complex organic sulfur compounds, suchas sulfides, disulfides, benzothiophene and the like, from a mixedhydrocarbon feedstock by catalytic reaction with hydrogen to formhydrogen sulfide.

Similar, and often simultaneously to hydrodesulfurization ishydrodenitrogenation, where complex organic nitrogen components areconverted with hydrogen to form hydrocarbons and ammonia. Typicalorganic nitrogen components are pyrrole, pyridine, amines andbenzonitriles.

Another refining application is hydrocracking which is used to reducethe boiling point of the feed by cracking large molecules into smallerones and adding hydrogen to them using a bifunctional catalyst.

Catalytic hydrotreating is, in all the above cases, a heterogeneousprocess, typically operated as a vapour phase process or as a liquid/gasphase process. In the conventional multi-stage hydrogenation processesthe hydrogen-containing gas and the material to be hydrogenated are fedthrough the plant in co-current or in counter-current fashion. In orderto achieve good economy of hydrogen usage, sometimes a recycle gas isused, typically comprising H₂ and a diluent such as methane other lightproduct gases of the main process.

The term “trickle bed reactor” or “trickle bed state” is often used todescribe a reactor in which a liquid phase and a gas phase flowcocurrently downward through a fixed bed of catalyst particles whilereaction takes place. However, these reactors can be operated in variousflow regimes, depending on vapor and liquid flow rates and properties.At sufficiently low liquid and gas flow rates the liquid trickles overthe packing in essentially a laminar film or in rivulets, and the gasflows continuously through the voids in the bed. This is termed the gascontinuous region or more specific “trickle flow regime” and is the typeencountered usually in refinery applications, in which typically largeexcess of hydrogen is required to prevent coking and to keep theconcentration of catalyst poison such as hydrogen sulfide that is formedduring the reaction low. It should be noted, however, that the operatingwindow of trickle flow is very wide and not only determined by flowrates (see, e.g., E. Talmor, AIChE Journal, Vol. 23, No. 6, 868-874(November 1977) discussed more fully below). Thus, for instance, andwithout wishing to be bound by theory, it may be possible to operatewith low liquid rates but at relatively high gas rates, too.

As gas and/or liquid flow rates are increased there is encounteredbehavior described as rippling, slugging, known in the art as “pulseflow regime”. Such behavior may be characteristic of the higheroperating rates often encountered in commercial petroleum processing.Pulsing is caused by alternating zones that are rich in vapor or inliquid. It is often called “high interaction flow regime”. At highliquid rates and sufficiently low gas rates, the liquid phase becomescontinuous and the gas passes in the form of bubbles; this is termed“dispersed bubble flow” or “bubble flow regime” and is characteristic ofsome chemical processing in which liquid flow rates (per unit crosssection area of the reactor) may be comparable to but more typically aremuch higher than the highest encountered in petroleum processing, butwhere gas/liquid ratios are much less.

Fixed bed hydrogenation reactors running in the bubble flow regime,i.e., relatively low volumetric gas to liquid ratio can have a tendencyfor the phases to separate. Without wishing to be bound by theory, suchphase separation may happen because the separate flows have a pressuredrop over the reactor or part of the reactor that is lower than themixed phase pressure drop would be; notionally, this may result in lossof much of the reaction due to starvation of one component in thosevapor-liquid separated zones, “hot-spots” resulting from the localizedaccumulation of reaction heat that is not transported away by anadequate flow of liquid, and/or overall axial or centerline temperatureprofile showing a higher ΔT than what is theoretically possible underideal flow conditions.

The aforementioned problem is avoided in most refinery processes where alarge excess of gas is recycled over the reactor, putting it into thetrickle flow or pulsing regime. Another way to solve the problem is touse high liquid velocities, with resulting high pressure drop along withhigh energy consumption. Both alternatives are high cost solutions.

What is needed is a simple and direct way of optimizing reactoroperation and/or design to combine the key variables such as pressure,feed flows, particle size, bed void fraction, and reactor dimensions,and to do so economically. Unfortunately, the literature appears to beconsistent in describing the phenomena of multiphase flow in a packedbed as complex and not well understood. Once the process chemistry,catalyst, and temperature are defined, the engineer is faced with theselection of a number of parameters which will determine the overalleffectiveness of the process. Variables include reactor diameter,reactor length, pressure, liquid rate, gas rate, gas/liquidstoichiometry, catalyst particle size, catalyst particle shape, catalystloading density, number of vessels, heat removal method, among others.

Undaunted by the complexity of the system, there have been numerousattempts in the past to optimize reactor operation and/or design usingselected variables. Some limited successes have been claimed for certaintypes of chemical reactions and/or using certain flow regimes.

U.S. Pat. No. 4,288,640 describes cocurrently passing a gas and liquidthrough a packed column in the form of a turbulent stream at rates suchthat, when the gas flow rate is kept constant, a variation in flow rateof the liquid produces a rise in the pressure difference Δp withincreasing liquid flow load L, expressed as Δp/L, at least twice aslarge as the rise Δp/L under liquid trickling conditions providinglaminar flow of the liquid over the packing bodies and continuous gasflow but the liquid flow rate being below the rate at which pulsing Δpin the column is produced, using specified packing.

U.S. Pat. No. 5,081,321 describes a catalytic hydrogenation process toproduce isopropanol by feeding hydrogen gas and acetone into a fixed bedreactor forming a cocurrent gas-liquid downflow while maintaining thecatalyst bed in a trickle bed state, wherein the following equation ismet: {B/[A·(σ/100)]}>1, wherein B is moles of hydrogen, A is moles ofacetone, and σ is percent conversion of acetone, provided that a trickleflow state is maintained.

U.S. Pat. No. 5,093,535 teaches a cocurrent hydrogenation process havinga hydrogenation zone containing a bed of catalyst whose particles lie inthe range of 0.5 mm to 5 mm and maintaining the supply of feedstock tothe bed so as to maintain a superficial liquid velocity of liquid downthe bed in the range of from about 1.5 cm/sec to about 5 cm/sec whilecontrolling the rate of supply of the H₂-containing gas to the bed so asto maintain at the top surface of the bed of catalyst particles a flowof H₂-containing gas containing from 1.00 to about 1.15 times thestoichiometric quantity of H₂ theoretically necessary to convert theorganic feedstock to the hydrogenation product.

FR 2,597,113 relates to trickle phase method for the selectivehydrogenation of highly unsaturated hydrocarbons, the methodcharacterized in that the product used, which contains the highlyunsaturated components and the hydrogenation gas, which containshydrogen, are directed through the catalyst at a surface flow velocitywith respect to the geometric surface of the particles of the totalquantity of catalyst of 1.5×10⁻⁷ to 3.0×10⁻⁵ m/s relative to thehydrogenation phase.

The prior art methods, however, suffer inter alia by making use of onlya very small number of reactor variables available and thus are toorestrictive, and/or do not provide results on pilot plant or laboratoryscale operations that can consistently be scaled up to commercial scalereactors.

E. Talmor, AIChE Journal, Vol. 23, No. 6, p. 868-874 (November 1977)teaches one type of flow map to describe the flow regimes in downflowthrough packed beds and found that the flow regimes encountered in suchsystems depend on the superficial volumetric gas-to-liquid ratio and theratio of inertia plus gravity forces to viscous plus interphase forces.

Flow maps attempt to predict the flow regime at a given set ofmeasurable conditions, but studies of this sort do not usually makepredictions or recommendations as to how well a chemical reaction wouldoccur in the predicted flow regime.

Flow maps have been used in a number of patents to ascertain usefulhydraulic regimes (e.g., trickle, pulse, or bubble), e.g., U.S. Pat. No.6,774,275 (WO2004016714) and the aforementioned U.S. Pat. No. 5,081,321.Numerous other publications concern this subject matter, e.g., Shermanet al., “Kinetic and Hydrodynamic Effects in the Activity Testing ofHydrodesulfurization Catalyst in Packed-Bed Reactors”, Symposium,Modeling and Troubleshooting of Commercial-Scale Reaction Systems, AIChE71st Annual Meeting, Nov. 12-16, 1978; Morsi et al., “Flow Patterns andSome Holdup Experimental Data in Trickle-Bed Reactors for Foaming,Nonfoaming, and viscous Organic Liquids”, AIChE Journal, Vol. 24, No. 2,pp. 357-360, March 1978. Talmor-like maps do not per se concern optimalconditions for a fixed bed reactor; they attempt to predict the flowregime. Typically flow maps are designed using water and air and failwhen applied to other fluid systems.

Among other references discussing hydraulic conditions or relatedfactors for operating a reactor include: U.S. Pat. Nos. 4,851,107;6,492,564; and 6,680,414; Holub et al., “Pressure Drop, Liquid Holdup,and Flow Regime Transition in Trickle Flow, AIChE Journal, Vol. 39, No.2, pp. 302-321, February 1993; Tosun, “A Study of Cocurrent Downflow ofNonfoaming Gas-Liquid Systems in a Packed Bed. 1. Flow Regimes: Searchfor a Generalized Flow Map”, Ind. Eng. Chem. Process Des. Dev. 1984, 23,29-35; Cheng et al., “Influence of hydrodynamic parameters onperformance of a multiphase fixed-bed reactor under phase transition”,Chemical Engineering Science 57 (2002) 3407-3413; Stuber et al.,“Partial Hydrogenation in an Upflow Fixed-Bed Reactor: A multistageOperation for Experimental Optimization of Selectivity”, Ind. Eng. Chem.Res. 2003, 42, 6-13; Al-Dahhan et al., “High Pressure Trickle-BedReactors: A Review, Ind. Eng. Chem. Res. 1997, 36, 3292-3314; Moreira etal., “Influence of Gas and Liquid Flow Rates and the Size and Shape ofparticles on the Regime Flow Maps Obtained in Cocurrent Gas-LiquidDownflow and Upflow through Packed Beds”, Ind. Eng. Chem. Res. 2003, 42,929-936; Attou et al., “A two-fluid hydrodynamic model for thetransition between trickle and pulse flow in a cocurrent gas-liquidpacked-bed reactor”, Chemical Engineering Science 55 (2000) 491-511;Herskowitz et al., “Effectiveness Factors and Mass Transfer inTrickle-Bed Reactors”, AIChE Journal, Vol. 25, No. 2, pp. 272-283;Gianetto et al., “Hydrodynamics and Solid-Liquid ContactingEffectiveness in Trickle-Bed Reactors”, AIChE Journal, Vol. 24, No. 6,pp. 1087-1104; Worstell et al., “Properly Size Fixed-Bed CatalyticReactors”, Chemical Engineering Progress, June 1993, pp. 31-37;Dudukovic, “Catalyst Effectiveness Factor and Contacting Efficiency inTrickle-Bed Reactors”, AIChE Journal, Vol. 23, No. 6, pp. 940-944;Morita et al., “Mass Transfer and Contacting Efficiency in a Trickle-BedReactor”, Ind. Eng. Chem, Fundam., Vol. 17, No. 2, 1978; Ng et al.,“Trickle-Bed Reactors”, Chemical Engineering Progress, November 1987,pp. 55-70; Borkink et al., “Influence of Tube and Particle Diameter onHeat Transport in Packed Beds”, AIChE Journal, Vol. 38, No. 5, May 1992,pp. 703-715; Talmor, “Part II. Pulsing Regime Pressure Drop”, AIChEJournal, Vol. 23, No. 6, pp. 874-878; Sai et al., “Pressure Drop inGas-Liquid Downflow Through Packed Beds”, AIChE Journal, Vol. 33, No.12, December 1987, pp. 2027-2036; Borio et al., “Cocurrently-CooledFixed-Bed Reactors: A Simple Approach to Optimal Cooling Design”, AIChEJournal, Vol. 35, No. 11, pp. 1899-1902, November 1989; Satterfield etal. “Mass Transfer Limitations in a Trickle-Bed Reactor”, AIChE Journal,pp. 226-234, March 1969; Charpentier et al., “Some Liquid HoldupExperimental Data in Trickle-Bed Reactors for Foaming and NonfoamingHydrocarbons”, AIChE Journal, Vol. 21, No. 6, pp. 1213-1218, November1975; and Burghardt et al., Chemical Engineering Science 57 (2002)4855-4863.

The present inventors have surprisingly discovered a method of reducingthe multidimensional problem outlined above to two dimensions.

SUMMARY OF THE INVENTION

The invention relates to a method whereby the numerous variables used tocarry out a reaction (or even mass transfer) in a fixed bed reactor canbe reduced to, or summarized by, two variables: Ta and φ. In anembodiment, Ta and φ are used as coordinates. Plotting pluralcoordinates (Ta, φ) based on a variety of conditions, a graph isobtained from which boundary conditions defining operating conditionsfor a desired process, such as a chemical reaction or mass transfer, maybe determined. In another embodiment, coordinates (Ta, φ) are determinedfrom a complete set of operating parameters and a determination is madewhether or not the operating parameters fall within desired boundaryconditions.

The present inventors have also surprisingly discovered an optimalhydraulic regime for conducting three-phase reactions or mass transferin a fixed bed reactor. In a preferred embodiment, the defined hydraulicregime provides a convenient method for design, operation, and/ordiagnosis of an hydrogenation process in a fixed bed reactor.

Included within the term “reaction”, for the purposes of the presentinvention, is mass transfer. While the use of a fixed bed comprising acatalyst is contemplated in preferred embodiments, the present inventionis applicable also to reactors having a fixed bed of any particulate orgranular material which is intended to interact with a liquid and/orgaseous material flowing through said reactor.

The invention is directed, in another embodiment, to the determinationof appropriate hydraulic conditions for a fixed bed reactor from a plotof plural (Ta, φ) coordinates, wherein Ta is defined as the sum of theinertia and gravity forces at a preselected point in said reactordivided by the sum of the interface and viscous forces at saidpreselected point in said reactor, and φ (“phi”) is defined as the gasto liquid volumetric flow ratio at said preselected point in saidreactor.

In still another embodiment, a plot of plural (Ta, φ) coordinates isused to provide a reactor or system of fixed bed reactors that meet therequirements of the desired conversion of reactants, supplying anappropriate amount of hydrogen, removing the heat of reaction to controlthe temperature within the desired range, and ensuring that thehydraulic conditions at a plurality and preferably substantially all ofthe points in the fixed bed are optimized such that, in a preferredembodiment, the flow is radially uniform and the gas is uniformlydispersed as fine bubbles throughout the length of the reactor.

In a further embodiment, the substantially optimum hydraulics aredetermined by setting the design variables, such as volumetric gas toliquid ratio, according to the criteria given herein.

In another embodiment, the invention is directed to a fixed bed reactoroperating under the following hydraulic conditions:Ta<N; and   (a)φ>a+(b·Ta);   (b)wherein variables N, a, and b are predetermined for a given reactionfrom the plot of plural (Ta, φ) coordinates, and Ta is further definedby the following equation:Ta=(1+1/Fr)/(We+1/Re);   (c)wherein Fr is the Froude number, We is the Weber number, and Re is theReynolds number. The Froude, Weber, and Reynolds numbers are per seknown in the art.

In yet another embodiment, the invention is directed to the designand/or operation of fixed-bed, gas-liquid, downflow hydrotreatingprocess, preferably a hydrogenation process, and preferably operatingunder one or more of the aforementioned embodiments. In a preferredembodiment the hydrogenation process comprises the hydrogenation ofcarbonyl moieties, more preferably hydrocarbon species having aldehydeand/or ketone.

It is an object of the invention to provide a method or system fordesigning, operating, monitoring and/or diagnosing a fixed bed reactionor reactor, particularly a hydrogenation or hydrotreating process or anymass transfer process using a fixed bed reactor.

It is yet another object of the invention to define conditions where thereactor hydraulics allow for maintenance of good gas/liquid contact andmass transfer and thereby provide for more efficient use of hydrogenand/or avoiding the need to recycle large amounts of hydrogen.

These and other embodiments, objects, features, and advantages willbecome apparent as reference is made to the following figures, detaileddescription, preferred embodiments, examples, and appended claims.

BRIEF DESCRIPTION OF THE DRAWINGS

In the accompanying drawings, like reference numerals are used to denotelike parts throughout the several views.

FIG. 1 illustrates plural (Ta, φ) coordinates taken at the inlet andoutlet within a reactor in an embodiment of the invention.

FIG. 2 illustrates the effect of changing a reactor variable for anembodiment of the invention.

FIG. 3 illustrates a plot of Ta versus (p for an embodiment of theinvention.

FIGS. 4, 5, and 6 illustrate the effect of a temperature pulse atvarious points in a reactor for an embodiment of the invention.

FIG. 7 illustrates a radial temperature profile for an embodiment of theinvention.

DETAILED DESCRIPTION

The present invention may be characterized, in an embodiment, as areduction of a multidimensional problem to two dimensions: Ta and φ,wherein the variable Ta is defined herein as the sum of inertia andgravity forces divided by the sum of the interface and viscous forces,and the variable φ is defined herein as the volumetric gas to liquidratio. Appropriate reactor conditions for a fixed bed catalytic reactormay be determined from plural coordinates (Ta, φ), as described indetail below. In an embodiment, appropriate reactor conditions may bedetermined from an orthogonal (x, y) plot where, for instance, Ta isx-axis and φ is the y-axis.

Ta may be further characterized, in an embodiment, by the expressionTa=(1+1/Fr)/(We+1/Re). Each of these expressions will be described indetail below. φ may be further characterized, in an embodiment, by theexpression u_(G)/u_(L).

The terms in the expression (1+1/Fr)/(We+1/Re) have the meanings setforth in Talmor, AIChE 23, No. 6, pp. 868-873, discussed in theBackground section, above, i.e, Fr is the Froude number, We is the Webernumber, and Re, the Reynolds number, are terms of art havingwell-defined meanings as set forth in the aforementioned Talmor article,and which may be determined by one of ordinary skill in the art.

The parameters in the aforementioned expressions are defined as follows:Fr=[(L+G)υ_(LG)]² /gD _(h)We=D _(h)(L+G)²υ_(LG)/σ_(L)Re=D _(h)(L+G)/μ_(LG)υ_(LG)=(L/G)υ_(L)/(1+L/G)+υ_(G)/(1+L/G)μ_(LG)=(L/G)μ_(L)/(1+L/G)+μ_(G)/(1+L/G)D _(h)=2εD/[2+3(1−ε)(D/D _(p))]where

-   D=column diameter, m-   D_(h)=bed hydraulic diameter, m-   D_(p)=equivalent particle diameter of catalyst, m-   Fr=Froude number, unitless-   G=superficial mass velocity of gas, kg/m²s-   g=acceleration due to gravity, m/s²-   L=superficial mass velocity of liquid, kg/m²s-   Re=Reynolds number, unitless-   u_(G)=ideal superficial velocity of gas at reactor temperature and    pressure, m/s-   u_(L)□=superficial velocity of liquid, m/s-   We=Weber number, unitless-   ε=void fraction of catalyst bed, unitless-   μ_(G)=viscosity of gas, kg/ms-   μ_(L)=viscosity of liquid, kg/ms-   μ_(LG)=effective viscosity of the gas-liquid mixture, kg/ms-   ρ_(G)=ideal gas density at reactor temperature and pressure, kg/m³-   ρ_(L)=liquid density, kg/m³-   σ_(L)=liquid surface tension, N/m-   υ_(G)=ideal specific volume of gas at reactor temperature and    pressure, m³/kg,-   υ_(L)=specific volume of liquid, m³/kg-   υ_(LG)=specific volume of the gas-liquid mixture, m³/kg

For spherical catalyst shape, D_(p) is the diameter of the sphere, andfor non-spherical catalyst geometry, the equivalent particle diameter iscalculated from the volume of the particle in cubic meters, V_(p), andthe external surface area of the particle in square meters, A_(p),according to the equation D_(p)=6V_(p)/A_(p).

In an embodiment, appropriate reactor conditions for a fixed bedcatalytic reactor may be determined from a plot of plural coordinates(Ta, φ). By way of example, a chemical process is carried out underpredetermined conditions in fixed bed catalytic reactor and the valuesTa and φ for each predetermined condition is plotted on a graph ascoordinates (x, y)=(Ta, φ). It will be understood by one of ordinaryskill in the art that the coordinates may also be plotted in numerousways, such as (x, y)=(φ, Ta), using a three-dimensional plot by addinganother variable, and so on. It will be further understood by the sameartisan that a physical plot is not necessary, but rather a computerprogram may be written to determine appropriate boundary conditionsand/or whether or not a given set of coordinates (Ta, φ) fall withinpredetermined boundary conditions.

The term “predetermined” as used herein means “determined or selectedbeforehand”.

In an embodiment, an indication is made on the graph regarding whetherthe results obtained are satisfactory or not. Whether or not conditionsare satisfactory or not can be determined by one of ordinary skill inthe art. It may be a subjective determination or an objectivedetermination or a combination thereof. Acceptable operating conditionsmay mean, by way of example, that part of a reactor is operable and notsubject to hotspots or runaway, and some larger part of the reactor isoperating in the optimal regime.

Conditions are varied so that a number of results are obtained, whichmay include both satisfactory and unsatisfactory experimental results,so that a plot of plural (Ta, φ) coordinates for a given chemicalreaction within a given reactor (or multiple reactors in series,parallel, or a combination thereof) having a fixed bed. In anembodiment, variables for changing conditions include one or more of thefollowing: reactor temperature, reactor diameter, reactor length,pressure, liquid rate, gas rate, gas/liquid stoichiometry, catalystparticle size, catalyst particle shape, catalyst loading density,catalyst bed void fraction, number of vessels, presence or absence ofrecycle and recycle stoichiometery (e.g., presence or absence of H₂ inrecycle, presence or absence of product recycle), and heat removalmethod. Variations on heat removal method include: adiabatic, recyclewith external cooling, recycle injected in one or more points of thereactor, coils or tubes within the catalyst bed, and mixtures thereof.

The coordinates (Ta, φ) may all be determined at the same point withinthe reactor, e.g., at the reactor inlet for cocurrent flow, but morepreferably coordinates (Ta, φ) are determined for numerous points withinthe reactor, e.g, reactor inlet and reactor outlet and at one or morepoints in the bed, in order to optimize the reactors performance inproducing product.

A boundary line that differentiates good from bad results will becomeapparent from the plot of plural coordinates (Ta, φ) to one of ordinaryskill in the art. Thus, it has surprisingly been found that pluralcoordinates (Ta, φ) map out an orthogonal space, with “good” or optimumpoints falling in a distinct region and “bad” or inoperable pointsfalling in another distinct region, with “marginal” or operable pointsfalling in yet another separate region between the optimal andinoperable regions. The user may select more refined differentiation ofthe coordinates, e.g., optimum and suboptimum, or good, operable, notoperable, and so on. Such characterization is fully within the skill ofthe ordinary artisan in possession of the present disclosure. Theboundary line(s) defining desired operating regions may be determinedqualitatively or it may be determined quantitatively, such as by adetermination of a boundary defined by a slope, y-intercept, andappropriate operating conditions may include other boundary conditions,as discussed below in the examples. Preferred operating conditions for ahydrogenation reaction in a fixed bed catalytic reactor according to thepresent invention include that the hydraulic conditions at a pluralityand preferably substantially all of the points in the fixed bed areoptimized such that the flow is radially uniform and the gas isuniformly dispersed as fine bubbles throughout the length of thereactor.

Boundary conditions may be assigned to preferred operating conditions.By way of example, according to an embodiment of the invention, a fixedbed reactor is operated under the following hydraulic conditions: (a)Ta<N; and (b) φ>a+(b·Ta), wherein, in addition to Ta and φ, thevariables N, a, and b are predetermined for a given reaction. Theparameters N, a, and b are predetermined from a plot of φ versus Ta, asstated previously. As will be apparent to one of ordinary skill in theart in possession of this disclosure, in the equation φ>a+(b·Ta), theterm “a” is the y-intercept and the term “b” is the slope for a boundaryline drawn, as shown in FIGS. 1 and 3, and in the equation Ta<N, theterm “N” is the limit of Ta determined by the boundary line drawn inthese same figures. It will also be apparent to the person of ordinaryskill in the art reading this disclosure that more complicated boundarylines and thus more complicated equations may be developed, but one ofthe advantages of the present invention is its simplification of acomplex problem. It will also be apparent to one of ordinary skill inthe art in possession of the present disclosure that a physical plot isnot necessary but rather a computer program could be written and used todefine a set of fully defined boundary conditions using the parametersTa and φ as set forth herein.

In an embodiment, Ta is less than about 600 or more preferably less thanabout 500. In another embodiment, Ta is greater than about 10. In yetanother embodiment, Ta is greater than about 10 and less than about 500,or less than 600.

In a preferred embodiment φ<0.8. In another preferred embodiment φ<1.1.In yet another preferred embodiment, the only limit on φ is that thereactor operate in the bubble flow regime, which may be determined byroutine experimentation by one of ordinary skill in the art inpossession of the present disclosure.

In another preferred embodiment, the reaction conditions are maintainedaccording to the following hydraulic conditions:10<Ta<500; and   (a){0.045+(0.00035·Ta)}<φ<0.8   (b)

In an embodiment, reactor hydraulic conditions, expressed by one or moreof the above embodiments, will hold in a plurality of points in saidreactor, preferably wherein conditions are sufficient to allow for achemical reaction to occur, and more preferably in the entirety of saidreactor.

In an embodiment, the conditions in a reactor may be monitored andmaintained under the aforementioned hydraulic conditions, so that thedesired product is obtained and flow is radially uniform and the gas isuniformly dispersed as fine bubbles throughout the length of thereactor.

In an embodiment, if conditions drift or otherwise are determined to beoutside of the hydraulic conditions according to the invention, then atleast one of temperature or pressure should be modified so as toobtained the desired hydraulic conditions. If this does not provide thedesired results, then one of more of the following should be modified:changing the composition of at least one of the gaseous or liquid feeds(e.g., adding diluents such as inert gases or liquids, partiallyrecycling product, adding more reactant gas and/or liquid, addingsurfactants, and the like), catalyst bed particle size or shape, orcatalyst bed void fraction, or diameter, length and/or number of beds,or a combination of these parameters. It will be recognized that atemperature range, pressure range, and catalyst composition are usuallydetermined based on chemistry and pilot plant data prior to the designof the commercial reactor, which places a practical limit on thepreferred options.

Pressure (or gas density) and temperature remain the key or preferredhydraulic parameters that may be adjusted in meeting the conditions ofthe present invention. The other preferred hydraulic variables arecatalyst particle size and shape, liquid velocity, gas velocity, and bedvoid fraction.

The effects of bed void fraction may be conveniently determined byextrapolation based on first principles and visual observations in atransparent laboratory reactor using water and air. Combinations of twoor more of the preferred hydraulic variables—catalyst particle size,liquid velocity, gas velocity, pressure, or a combination thereof—arealso preferred embodiments of the present invention.

In a preferred embodiment the reaction is a hydrotreating proceess,preferably a hydrogenation process, more preferably comprising thehydrogenation of carbonyl moieties, still more preferably thehydrogenation of hydrocarbon species having aldehyde and/or ketonemoieties and yet still more preferably the hydrogenation processcomprises the hydrogenation of aldehydes and/or ketones produced in theOxo Process, more yet again more preferably such a process using an Oxohydrogenation catalyst, said catalyst in the form of extrudates which,in one embodiment, have from about 1 to about 4 mm nominal diameter witha length to diameter ratio of about 3 or more, in another embodimenthave a trilobed or quadralobed shape, and in a still yet more preferredembodiment have the aforementioned diameter and shape.

The invention is not specific to any particular reaction or catalyst. Itis specific, however, to fixed bed reactors having three phases: thesolid phase comprising a fixed bed catalyst, a liquid phase, and a gasphase. Thus, the actual chemical reaction occurring within the systemmay be, for instance, an hydrogenation reaction, a hydrodesulfurizationreaction, a water treatment process, and the like. The invention isparticularly beneficial in the design, operation, and diagnosis of athree-phase reaction in a fixed bed reactor having a catalyst therein,providing a convenient tool to increase gas-liquid-catalyst interactionover the full length and diameter of the reactor.

In a more preferred embodiment wherein the chemical reaction is anhydrogenation process comprising the hydrogenation of aldehydes and/orketones to alcohols, the process of the invention preferably operatesalways in the dispersed bubble flow regime in a cocurrent downflowscheme. In regimes other than cocurrent downflow operating undersubstantially dispersed bubble flow regime, the axis of the map providedby Ta versus φ may need to be modified to account for different forcesacting on the fluids, as would be recognized by one of ordinary skill inthe art in possession of the present invention (bearing in mind thatnumber Ta has been previously defined broadly as the inertia and gravityforces divided by the interface and viscous forces). Nevertheless thepresent invention provides a good starting point for optimizingconditions under different hydrodynamics than this preferred embodiment.

Thus, in a preferred embodiment, the hydrogenation process is operatedunder conditions sufficient to provide for bubble flow regime.

Various embodiments as set forth herein, preferred and otherwise, may becombined as would be readily apparent to one of ordinary skill in theart in possession of the present disclosure. Thus, for instance, in afurther embodiment, which is a preferred embodiment, the presentinvention may be applied to a cocurrent downflow hydrogenation processunder conditions sufficient to provide for dispersed bubble flow, andfurther characterized by the following conditions:10<Ta<500, and   (a){0.045+(0.00035×Ta)}<φ  (b)and in a yet more preferred embodiment wherein φ is less than 1.1, orless than 0.8.

In a more preferred embodiment, the aforementioned hydraulic conditionsare met in at least one three-phase reaction zone wherein thehydrogenation reaction is the hydrogenation of hydrocarbons selectedfrom C3-C20 hydrocarbons and mixtures thereof, wherein at least some ofsaid hydrocarbons have reactive carbonyl moieties. In an embodiment ofthis more preferred embodiment, the solid phase comprises the fixed bedcatalyst which includes a hydrogenation catalyst, e.g., a catalystselected from cobalt supported on alumina, molybdenum (Mo) supported onalumina, cobalt and molybdenum supported on alumina, nickel andmolybdenum supported on alumina, all of the above either reduced orsulfided, and Cu, Cr, Zn, and mixtures of aforementioned catalysts. Theliquid phase comprises the C3-C20 hydrocarbons having reactive carbonylmoieties, e.g., aldehydes and/or ketones, and the gas phase comprises asource of hydrogen, e.g., H₂. In another embodiment, there may bemultiple reactors, which may be a series of reactors or reactors inparallel, or a combination thereof, with the catalyst for each bedindependently selected. Also, even more preferably, there may bedifferent layers in the bed, each layer being of a different particlesize catalyst and/or loading density. This is particularly useful whengas depletion occurs due to reaction, and the flow conditions need to beadjusted.

In the aforementioned embodiment involving multiple reactors, it iscontemplated that either a single graph comprising plural plots of (Ta,φ) coordinates may be prepared for the entire system of plural reactorsor plural graphs of plural plots of (Ta, φ) may be prepared, forinstance, with a separate graph for each separate reactor.

In an embodiment of the above-recited preferred embodiment, the liquidphase comprises at least one aldehyde and/or ketone selected from C3-C20aldehydes and/or ketones, preferably C6-C14 aldehydes and/or ketones,and more preferably C8-C13 aldehydes or ketones. Typically the liquidphase feed will be a mixture of aldehydes and/or ketones having a rangeof carbon numbers and/or a isomers within a single carbon number. Astill more preferred embodiment of each of the aforementionedembodiments, the particular range of aldehydes and/or ketones comprisesaldehydes and/or ketones made in the Oxo aldehyde process, whereinolefins are reacted under hydroformylation conditions to yield thecorresponding aldehyde having one more carbon atom, as discussed in theBackground section.

As previously mentioned, the present invention may also be used inreactor design, operation, and/or diagnosis. By way of example, thehydrogen rate (or in the more general case, the gas velocity) may beselected at the desired stoichiometry and then the reactor diameter isselected to provide the preferred liquid velocity within a predeterminedgraph comprising plural (Ta, φ) coordinates. An option in the designand/or operation may then be selected to include recycling partiallyconverted product to make up part of the liquid feed.

By way of further example, a plot of (Ta, φ) coordinates may be used inthe diagnosis or selection of operating conditions, as illustrated inFIG. 1. In FIG. 1, each string of points represents plural coordinatesplotted for (x, y)=(Ta, φ) which are measured at the reactor inlet andoutlet (top and bottom points, respectively) and predicted forintermediate points throughout the reactor. In this example, allconditions are kept constant except catalyst particle size and gas flow,i.e., effecting D_(p) (equivalent particle diameter of catalyst), G(superficial mass velocity of gas), and u_(G) (superficial velocity ofgas), in the equations discussed above. Coordinates representinginoperable conditions are not shown on the graph. The solid line in FIG.1 representing the boundary conditions, is drawn through each string(and numerous other strings and points not shown on the graph) atapproximately the point where performance is judged to be marginal.These points within a reactor where performance deviates from theacceptable conditions may be determined by various tests, such as thosediscussed herein. Appropriate operating conditions are represented onthe graph by any point to the upper left of the solid line representingthe boundary conditions. FIG. 1 is intended to illustrate the inventionschematically; the actual numbers associated with the graph are notgermane to the example.

It should be noted that the present inventors have simplified the methodby using the ideal gas phase density PG rather than try to use theactual gas density inside the reactor, which would require takingaccount of the flashing of volatile components. This is a minorcorrection which simplifies the effort. In addition, by taking a sampleof liquid in and out of the reactor, the amount of H₂ consumed at anylongitudinal point in the reactor may be determined, and thus the phaseratio at any longitudinal point may be calculated. Likewise, in thepreferred embodiment illustrated in FIG. 1, different size particle ofcatalyst (D_(p)) and void fraction (υ) were used, and of course accountis taken of this fact for every longitudinal point studied in thereactor.

In an embodiment of the invention, starting point values of physicalproperties for a process comprising hydrogenation of an organicfeedstock in a downflow, cocurrent, fixed bed used in computationsherein are as follows: μ_(G)=1.6×10⁻⁵ kg/m−s; μ_(L)=2.9×10⁻⁴; ρ_(G)=12.3kg/m³; ρ_(L)=781 kg/m³; and σ_(L)=0.015 N/m. For such a system, valuesof various parameters will typically vary as follows: column diameter(D) from 0.4 to 2.0 m; particle diameter (D_(p)) from 1 to 3 mm;superficial liquid velocity u_(L) from 0.005 to 0.04 m/s; superficialideal gas velocity u_(G) from 0.001 to 0.01 m/s; and catalyst bed voidfraction (ε) from 0.38 to 0.54.

While in an embodiment optimal operating conditions specified by thepresent invention will exist in all parts of the reactor, from theinlet, through all levels in the catalyst bed, to the outlet, in otherembodiments only a portion of the reactor will have the specifiedhydraulic conditions, i.e., a plurality of points in said reactor. Onemay, for example, choose to have part of a reactor operating in a regimethat is merely “operable” but not subject to hotspots or runaway andsome larger part of the reactor operating in the “optimal” regime.

The present invention has been described generally above with referenceto certain embodiments. The following specific examples are provided asrepresentative examples and are not intended to limit the invention.

EXAMPLES

The following examples describe experiments using commercial reactors inOxo aldehyde hydrogenation service. The examples illustrate how thepresent invention may be used to select a range of preferred and/oroptimal conditions for a selected reaction.

Testing was done on three commercial reactors. The effects on reactorperformance of pressure, gas rate, liquid rate, and particle size weremeasured by static and dynamic test methods. The performance wasmeasured by analysis of sample, overall heat balance, radial temperatureprofiles, axial temperature profiles, and dynamic temperature responseto temperature pulses, and/or temperature steps under both reacting andnon-reacting conditions.

The results of each test were evaluated based on at least two of theabove diagnostic methods, and were judged qualitatively to be one ofthree possible categories: (1) good—the reactor performs as it wasdesigned, is predictable and stable; (2) operable—the reactor can be runsafely and continuously, but it is not performing up to its design basisbased on low activity, under-utilization of the full length of the bed,poor dynamics, or some other reason; or (3) not operable—some part ofthe reactor is not responsive or is highly sensitive to one or more ofthe manipulated variable that an operator may use to set his operatingconditions, which may include a hot-spot, cold spot, and the like. Inreality there is a continuum of performance definitions across regimes(1) and (2), but they were not qualified further. The boundary linedrawn between the good points (1) and the operable points (2) thus showsthe boundary criterion for the effective design and operating conditionsfor fixed bed reactors operating as hydrogenation reactors for Oxoaldehyde hydrogenation. The inoperable points are left as a subset ofthe operable, or “just operable” points.

Each of three commercial Oxo aldehyde hydrogenation reactors were fittedwith multiple thermocouples arranged from the inlet, through the reactorbed, to the outlet. Reactors A, B, and C are fixed beds having a waterjacket for protection of vessel integrity in case of emergency. Thereactors are essentially adiabatic. Reactors A and B have the capabilityof recycling partially converted material from the reactor outlet.Reactor C does not have recycle capability. In all of the examples thatfollow, physical properties of gas and liquid are effectively constant.

Example 1

Gas and liquid rates were varied in an attempt to render inoperablereactors operable. The liquid rate to reactor A was increased by about30% by decreasing the aldehyde feed and increasing the amount ofrecycle, and the gas rate was reduced by about 75%, still at about a 60%molar excess to the aldehyde content of the fresh feed. Reactor A wasinoperable as the automatic shutdown system was activated due to hightemperature. The (Ta, φ) coordinates for this example are (462, 0.083).A similar experiment was run on reactor B. The gas rate was increasedand the liquid rate was reduced to such an extent that the hightemperature shut down system was activated. The (Ta, φ) coordinates forthis inoperable condition were (1216, 0.24).

Example 2

This example compares the performance of reactor C with reactor A, eachhaving different size catalyst particles. Reactor C is 63% of the crosssectional area of reactor A. The catalysts in each reactor is the same,a proprietary catalyst supported on commercially standard extrudedsupports. They differ only in particle size: Reactor C uses 1.1 mmequivalent diameter catalyst while Reactor A uses 2.1 mm equivalentdiameter catalyst. Both reactors are fitted with a water jacket wheretempered water, somewhat cooler than the average reactor temperature, iscirculating. The water jackets are used to protect the vessel fromundetected hotspots that could melt through the reactor wall. Thejackets do not remove a significant amount of reaction heat since thereactor diameters are relatively large. Under ideal flow conditions,there should be no significant radial temperature gradient in thereactors. If a significant radial temperature gradient is measured inthe bottom of the catalyst bed, it can be averaged over the total crosssectional area of the bed to give a prediction of the temperature of thetotal reactor effluent in the outlet pipe. If the predicted value andactual value are different, one can determine an appropriate radial massflow distribution, which, if used in the temperature profile averagingprocess, gives the actual temperature of the reactor effluent. Since thereactors are liquid full and operate in the dispersed bubble flowregime, the radial mass flow distribution is essentially the liquid flowdistribution. This is the method by which the quality of flowdistribution was measured in this and other of the examples below.

In Example 2, reactor A was run with a higher superficial liquid massrate than reactor C by 22%. Reactor A was run with a higher gassuperficial mass rate than reactor C by 38%. Both reactors weremaintained at essentially the same temperature and pressure; only thevelocity of the liquid and gas were varied. The hydrogen ratescorrespond in both reactors to at least 60% molar excess with respect tothe molar amount of aldehyde in the feed to each reactor.

The Talmor map coordinates (Ta, φ) for reactors A and C are (500, 0.42)and (400, 0.37), respectively.

At these conditions, the centerline bed outlet temperature of reactor Awas 18° C. higher than the temperature of the mixed reactor effluent inthe outlet pipe two meters downstream. The radial temperature differencebetween the bed centerline and the wall region was 35° C. The apparentactivity of the catalyst in reactor A, as measured by sample, was lessthan the design value. Reactor C showed a relatively flat radialtemperature profile and a low temperature difference between the bedbottom and the outlet pipe. Results of the temperature measurementsindicated that reactor A had a radial mass flow maldistribution suchthat the liquid flowing near the wall was 3 to 4 times higher than theflow in the center of the reactor. By the same method of analysis,reactor C had a relatively uniform flow distribution. The smallerparticle size improved the performance of reactor C even at lower flowrates per unit area of liquid and gas. The performance of reactor C wasclassified as “good” and the performance of reactor A was classified asonly “operable” based on the measurements at the bottom of the catalystbed.

Example 3

Commercial reactor C, an adiabatic reactor, fitted with a water jacket(for safety concerns) and containing 1.1 mm equivalent diametercatalyst, was tested at different liquid and gas flow rates, as shown inFIG. 2. The progression of the reaction down the catalyst bed causes anincrease in the axial temperature profile measured near the center ofthe bed. The relative uniformity of the flow distribution in the reactorwas determined by comparing the centerline temperature at the bottom ofthe bed with the mixed effluent temperature. The mixed effluenttemperature provides a measure of the total temperature conditions inthe entire reactor. Thus, as the difference between the centerlinetemperature at the bottom of the bed and the mixed effluent temperatureapproaches zero, the liquid and gas distribution in the reactor iscloser to being uniformly radially distributed, ensuring that there areno significant zones of low flow with the possibility of hot spots orhydrogen starvation. Uniform liquid flow is indicated when the ΔT is1-2° C.

The impact of both liquid and gas rate are shown in FIG. 2. The reactorperformance is considered to be “good” with a ΔT less than about 4° C.Each point is plotted as a (Ta, φ) coordinate in FIG. 3. This set ofexperiments indicates that the reactor has very poor packingdistribution which causes temperature maldistribution, and may bebrought into good operation by increasing the mass flow (one or both ofliquid rate and gas rate) and/or decreasing bed void fraction.Increasing mass flow would typically be adopted in the case where acommercial reactor was operating, whereas decreasing bed void fraction,and/or particle size, and/or changing reactor size typically might beadopted in the design of a reactor.

Example 4

This test illustrates an alternate or supplemental diagnostic method toconfirm that the hydrogenation process was operating in the desiredhydraulic regime. A pulse increase in the feed temperature wasintroduced into reactor A of Example 2 and at the same operatingconditions as Example 2. At arbitrary time 6.5, the steam valve on thefeed temperature conditioning control system was opened forapproximately one unit of time to allow the reactor feed temperature toincrease by 5° C. At that time, the temperature controller was put intoautomatic at the original setpoint. The temperature responses downseveral axial positions of the reactor were recorded by means of thereactor thermocouples and standard plant instrumentation, as shown inFIG. 4.

FIG. 4 shows the feed temperature, a catalyst bed temperature 1.5 metersfrom the inlet distributor, a temperature in the middle of the bed, anda temperature 1.5 meters from the outlet of the bed. The shapes anddelay times of the first two thermocouples are in line with what wouldbe expected from a catalyst bed with well distributed uniform flowacross the cross section of the reactor. Near the bottom of the bed,there is not enough flow over the last thermocouple to change itstemperature significantly. Radial temperature measurements showed thisreactor to have about 4 times the liquid velocity near the wall regionthan at the center-line. The consumption of gas in passing through thereactor has depleted the gas enough to make the bubble flow regimeunstable, allowing the gas and liquid to separate into distinct regions.The coordinates of this process condition in terms of (Ta, φ) are givenin Example 2.

Example 5

Dynamic testing of two reactors type A under recirculating/no reactioncondition are illustrated in FIGS. 5 and 6.

Reactor A was tested with circulation of fully converted liquid and nofresh feed. Clearly, there is no consumption of gas as the two phasestravel down the reactor. Initially the reactor was at 135° C. when, attime 5, the set point on the feed temperature was reduced by 30° C. Thetransient response of all the axial temperature measurements wererecorded. Two tests were done at constant liquid flow rate. FIG. 5illustrates the axial temperature response to step change in feedtemperature at high gas rate and FIG. 6 illustrates the axialtemperature response to step change in feed temperature at low gas rate.The gas rate in FIG. 5 is 3.5 times the gas rate in FIG. 6. At thehigher gas rate, it can be seen from FIG. 5, the temperature responsesfrom top to bottom of the reactor are nearly identical (and thus nearlyideal), indicating that the flow is near plug flow and the liquiddistribution remains uniform across the cross section of the reactor.

As shown in FIG. 6, the dynamic response of middle of the bedtemperature is similar to that of the higher gas rate example,indicating that the initial distribution of liquid and gas has beenmaintained for a certain distance into the reactor. The thermocouplenearest the bottom of the reactor does not show the sharp turndown ofthe previous test, and it falls much more slowly. This indicates thatthe flow near the center of the bed has slowed down considerably and hasmigrated toward the reactor wall, most likely with the gas migratingtoward the center. The bubble flow regime at these conditions of liquidand gas rate is not stable for the entire length of the reactor. Sincethere is no gas consumption, the gas to liquid ratio at all levels inthe reactor is the same. The time scales are identical for each of thecases. The (Ta, φ) coordinates of the high gas case are (272,0.32) andfor the low gas case are (462,0.08).

Example 6

According to Gupta (supra), a natural tendency for the phase separationexists because the gas and liquid phases flowing separately throughdifferent paths in the bed yield a lower pressure drop than when theyshare the same paths. The following example shows that when the gas issufficiently depleted by reaction, phase separation occurs. In thiscase, liquid migrated predominantly to the wall region and consequently,gas holdup increased in the center of the reactor. Example 2 showed lowflow in the centerline of the reactor that started below the midpoint ofthe bed.

At the same flow conditions at which the temperature pulse test was donein Example 4, a steady state radial temperature profile across thecatalyst bed was measured. When the temperature profile is averaged overthe cross sectional area of the reactor, appropriately weighted by themass flow rate at each radial position, or, more precisely, in each ofthe sequential annuli marked by the radial positions, the averagetemperature can be made to match the mixed effluent from the reactoroutlet temperature, assuming a radial mass flow profile. This assumedmass flow profile may be of any arbitrary shape, linear, parabolic, etc.It will be obvious to anyone skilled in the art that if the flow isradially uniform, the values of the flow profile computed to match theaverage bed temperature with that of the reactor outlet will berelatively constant across the radial direction of the reactor. Thesimplest profile to choose is a linear one.

The radial temperatures are compared with the outlet line temperature inFIG. 7. The linear velocity profile which matches the outlet linetemperature with the averaged radial temperatures shows that the massflowing near the wall to be several times what it is in the center ofthe reactor. The mass flow is essentially the liquid only flow. Thelower temperature near the wall is caused by the loss of adiabaticheating due to the starvation of hydrogen which has most certainlymigrated to the center of the reactor.

These examples illustrate the diagnostic methods, which when takentogether, unequivocally explain the quality of the three-phase bubbleflow regime over the entire length of the reactor, in terms ofeffectively conducting a hydrogenation reaction. Additional experimentswere performed at different conditions, each analyzed using one or moreof these diagnostic methods, to develop the optimum operating conditionsfor a hydrogenation process. These are shown in aggregate in FIG. 3. Forthis embodiment, boundary conditions for “good” operation are determinedto be 10<Ta<500 and {0.045+(0.00035·Ta)}<φ<1.1.

The present inventors have surprisingly found that in a preferredembodiment, process design decisions selected from at least one of:higher liquid velocity, higher gas velocity, smaller catalyst particlesize, lower bed void fraction, have the most positive effect on the safeand productive operation of a hydrogenation process. All these factorscome with a higher initial investment cost or higher operating cost. Ahydrogenation process designed and operated within the conditions of theTa and φ specifications of this invention does not require strictspecification of any parameter such as superficial velocity of gas orliquid, pressure drop, catalyst particle size, bed void fraction orreactor dimensions. Rather, it allows for the sensible tradeoff amongthese conditions to satisfy an optimization criteria that may bedifferent from one location to another, tradeoffs that would allow theuse of existing equipment such as reaction vessels, compressors, orpumps, special requirements dictated by the chemistry of the catalysisto be conducted, unavailability of a certain size of catalyst, catalystphysical properties such as crush strength, and the like.

The invention has been described above with reference to numerousembodiments and specific examples. Many variations will suggestthemselves to those skilled in this art in light of the above detaileddescription. All such obvious variations are within the full intendedscope of the appended claims. Particularly preferred embodimentsinclude: a method for determining appropriate hydraulic conditions for aprocess in a fixed bed reactor comprising plotting plural coordinates(Ta, φ) on a graph, each coordinate obtained by carrying out saidprocess in a fixed bed reactor under preselected hydraulic conditions,wherein Ta is the sum of the inertia and gravity forces divided by thesum of the interface and viscous forces in said reactor for said processat a preselected point in said reactor, and φ is the volumetric gas toliquid flow ratio in said reactor for said process at said preselectedpoint, and determining from said graph appropriate hydraulic conditions;or preferably further limited by at least one of the limitations setforth in the specification, which may be combined as would be apparentand practicable to one of ordinary skill in the art in possession of thepresent disclosure which may be combined as practicable, particularly:further including, after determining from said graph appropriatehydraulic conditions, carrying out said process in a fixed bed reactorunder said appropriate hydraulic conditions, and/or wherein said processin said fixed bed reactor is a hydrotreating process, and/or whereinsaid process in said fixed bed reactor comprises a downflow, cocurrent,fixed bed hydrogenation of an organic feedstock, and/or wherein saidprocess in said fixed bed reactor is running in a bubble flow regime,and/or wherein each coordinate (Ta, φ) is obtained under hydraulicconditions differing from the hydraulic conditions for all othercoordinates (Ta, φ) on said graph by at least one of: reactor diameter,reactor length, pressure, liquid rate, gas rate, gas/liquidstoichiometry, catalyst particle size, catalyst particle shape, catalystloading density, catalyst bed void fraction, presence or absence ofrecycle, recycle stoichiometry, number of vessels, and heat removalmethod, or wherein for each coordinate (Ta, φ), Ta is further defined bythe expression Ta=(1+1/Fr)/(We+1/Re), and φ is further defined by theexpression φ=u_(G)/u_(L), wherein u_(G) and u_(L) are the superficialflow velocity of gas and liquid phases in m/s, respectively, and furtherwherein:Fr=[(L+G)υ_(LG]) ² /gD _(h)We=D _(h)(L+G)²υ_(LG)/σ_(L)Re=D _(h)(L+G)/μ_(LG)υ_(LG)=(L/G)υ_(L)/(1+L/G)+υ_(G)/(1+L/G)μ_(LG)=(L/G)μ_(L)/(1+L/G)+μ_(G)/(1+L/G)D _(h)=2εD/[2+3(1−ε)(D/D _(p))]where

-   D=column diameter, m-   D_(h)=bed hydraulic diameter, m-   D_(p)=equivalent particle diameter of catalyst, m-   Fr=Froude number, unitless-   G=superficial mass velocity of gas, kg/m²s-   g=acceleration due to gravity, m/s²-   L=superficial mass velocity of liquid, kg/m²s-   Re=Reynolds number, unitless-   u_(G)=ideal superficial velocity of gas at reactor temperature and    pressure, m/s-   u_(L)□=superficial velocity of liquid, m/s-   We=Weber number, unitless-   ε=void fraction of catalyst bed, unitless-   μ_(G)=viscosity of gas, kg/ms-   μ_(L)=viscosity of liquid, kg/ms-   μ_(LG)=effective viscosity of the gas-liquid mixture, kg/ms-   ρ_(G)=ideal gas density at reactor temperature and pressure, kg/m³-   ρ_(L)=liquid density, kg/m³-   σ_(L)=liquid surface tension, N/m-   υ_(G)=ideal specific volume of gas at reactor temperature and    pressure, m³/kg,-   υ_(L)=specific volume of liquid, m³/kg-   υ_(LG)=specific volume of the gas-liquid mixture, m³/kg    and/or wherein said process further comprises passing a feed    solution comprising an organic feedstock downwardly in co-current    with a hydrogen-containing gas through a hydrogenation zone    comprising a bed of a hydrogenation catalyst, and wherein said    organic feedstock comprises at least one aldehyde or ketone selected    from C3-C13 aldehydes and ketones. Another particularly preferred    embodiment inlcudes: a process in a fixed bed catalytic reactor    operating under the following conditions:    Ta<N; and   (a)    φ>a+(b·Ta);   (b)    wherein variables N, a, and b are predetermined from a graph of    plural (Ta, φ) coordinates according to Claim 1, which may also be    further limited by the limitation that N, a, and b are predetermined    by drawing a line separating the coordinates (Ta, φ) providing    acceptable results from those coordinates (Ta, φ) providing    unacceptable results, said line having a y-intercept defined by the    parameter “a” and a slope defined by the parameter “b” in    equation (b) and having a limit of acceptable results defined by the    parameter “N” in said equation, and/or wherein said process    comprising passing a feed solution comprising an organic feedstock    downwardly in cocurrent with a hydrogen-containing gas through a    hydrogenation zone comprising a bed of a hydrogenation catalyst,    under hydraulic conditions that satisfy the following equations:    10<Ta<500; and   (a)    {0.045+(0.00035×Ta)}<φ<0.8.   (b)

Yet another preferred embodiment includes: a process comprising passinga feed solution comprising an organic feedstock downwardly in cocurrentwith a hydrogen-containing gas through a hydrogenation zone comprising abed of a hydrogenation catalyst, monitoring said process and takingmeasurements so as to calculate whether or not the following equationshold in at least one point in the reactor:Ta<500; and   (a)φ>a+(b·Ta);   (b)wherein variables a and b are predetermined, and wherein φ is the gas toliquid volumetric flow ratio, and Ta is the Talmor number defined by thefollowing equation:Ta=(1+1/Fr)/(We+1/Re);   (c)wherein Fr is the Froude number, We is the Weber number, and Re is theReynolds number, even more preferably but optionally further includingcalculating whether or not said equations hold for said reaction, andeven more preferably but optionally wherein if said equations do nothold for said reaction, changing at least one reactor variable selectedfrom reactor diameter, reactor length, pressure, liquid rate, gas rate,gas/liquid stoichiometry, catalyst particle size, catalyst particleshape, catalyst loading density, catalyst bed void fraction, presence orabsence of recycle, recycle stoichiometry, number of vessels, and heatremoval method, and yet still more preferably but optionally furtherincluding, after said changing of at least one reactor variable,repeating the steps of taking said measurements and calculating whetheror not said equations hold for said reaction. Yet still other preferredembodiments include an hydrogenation process in a downflowing fixed bedreactor comprising hydraulic conditions defined by the followingequations:10<Ta<500; and   (a){0.045+(0.00035×Ta)}<φ<0.8;   (b)wherein φ is the volumetric gas to liquid flow ratio and Ta is definedby the following equation:Ta=(1+1/Fr)/(We+1/Re);   (c)wherein Fr is the Froude number, We is the Weber number, and Re is theReynolds number; a method of producing an alcohol comprising feeding aliquid comprising an organic species comprising at least 3 carbon atomsand a reactive carbonyl moiety into a fixed bed reactor under conditionssufficient to cause hydrogenation of said reactive carbonyl moiety, saidconditions including hydraulic conditions determined according to Claim1; a process for carrying out a chemical reaction in a fixed bedreactor, the process comprising providing at least one zone having asolid catalyst, a liquid phase, and a gaseous phase, under conditionssufficient to carry out a chemical reaction, the improvement comprisingoperating said zone under hydraulic conditions obtainable according tothe method of Claim 1 (and a more preferred embodiment wherein saidhydraulic conditions are obtained by obtaining new (Ta, φ) coordinatesand comparing said new coordinates to the boundary conditions obtainedfrom said plot of (Ta, φ) coordinates, wherein said new coordinates areobtained by changing at least one of hydraulic diameter, reactordiameter, reactor length, pressure, liquid rate, gas rate, gas/liquidstoichiometery, catalyst particle size, catalyst particle shape,catalyst loading density, number of vessels, and heat removal method, oryet still more preferably wherein said process is repeated until (Ta, φ)coordinates are obtained wherein no free H₂ is observed at the end ofthe reactor, and even still more preferably wherein heat removal methodsare selected from adiabatic heat removal, heat removal by recycle withexternal cooling wherein the cooled recycle is injected in one or morepoints of the reactor, or by coils or tubes within the catalyst bed; andalso a process for carrying out a chemical reaction in a fixed bedreactor comprising determining the operating conditions required forsaid chemical reaction, carrying out said chemical reaction in saidfixed bed reactor, and obtaining at least one product from said chemicalreaction, the improvement comprising: plotting plural (Ta, φ)coordinates on a graph, each coordinate obtained by carrying out saidchemical reaction in said fixed bed reactor, wherein Ta is the inertiaand gravity forces divided by the interface and viscous forces at apreselected point in said reactor and φ is the volumetric gas to liquidratio at said preselected point, and obtaining from said graph boundaryconditions required for said operating conditions required for saidchemical reaction, and operating said chemical reaction within saidboundary conditions; and also a fixed bed catalytic reactor having atleast one of column diameter, column length, bed hydraulic diameter,equivalent particle diameter of catalyst, catalyst shape, and voidfraction of catalyst bed determined by a process according to theabove-recited processes; and also a graph comprising a plot prepared bythe above-recited processes; and also the use of said graph, such as fordetermining appropriate reactor conditions, to produce a product, todesign a reactor, to diagnose a reactor, or design a reactor part.

Trade names used herein are indicated by a ™ symbol or ® symbol,indicating that the names may be protected by certain trademark rights,e.g., they may be registered trademarks in various jurisdictions.

All patents and patent applications, test procedures (such as ASTMmethods, UL methods, and the like), and other documents cited herein arefully incorporated by reference to the extent such disclosure is notinconsistent with this invention and for all jurisdictions in which suchincorporation is permitted.

When numerical lower limits and numerical upper limits are listedherein, ranges from any lower limit to any upper limit are contemplated.While the illustrative embodiments of the invention have been describedwith particularity, it will be understood that various othermodifications will be apparent to and can be readily made by thoseskilled in the art without departing from the spirit and scope of theinvention. Accordingly, it is not intended that the scope of the claimsappended hereto be limited to the examples and descriptions set forthherein but rather that the claims be construed as encompassing all thefeatures of patentable novelty which reside in the present invention,including all features which would be treated as equivalents thereof bythose skilled in the art to which the invention pertains.

1. A method for determining appropriate hydraulic conditions for aprocess in a fixed bed reactor comprising plotting plural coordinates(Ta, φ) on a graph, each coordinate obtained by carrying out saidprocess in a fixed bed reactor under preselected hydraulic conditions,wherein Ta is the sum of the inertia and gravity forces divided by thesum of the interface and viscous forces in said reactor for said processat a preselected point in said reactor, and φ is the volumetric gas toliquid ratio in said reactor for said process at said preselected point,and determining from said graph appropriate hydraulic conditions.
 2. Themethod according to claim 1, further including, after determining fromsaid graph appropriate hydraulic conditions, carrying out said processin a fixed bed reactor under said appropriate hydraulic conditions. 3.The method according to claim 1, wherein said process in said fixed bedreactor is a hydrotreating process.
 4. The method according to claim 1,wherein said process in said fixed bed reactor comprises a downflow,cocurrent, fixed bed hydrogenation of an organic feedstock.
 5. Themethod according to claim 1, wherein said process in said fixed bedreactor is running in a bubble flow regime.
 6. The method according toclam 1, wherein each coordinate (Ta, φ) is obtained under hydraulicconditions differing from the hydraulic conditions for all othercoordinates (Ta, φ) on said graph by at least one of: reactor diameter,reactor length, pressure, liquid rate, gas rate, gas/liquidstoichiometry, catalyst particle size, catalyst particle shape, catalystloading density, catalyst bed void fraction, presence or absence ofrecycle, recycle stoichiometry, number of vessels, and heat removalmethod.
 7. The method according to claim 1, wherein for each coordinate(Ta, φ), Ta is further defined by the expression Ta=(1+1/Fr)/(We+1/Re),and φ is further defined by the expression φ=u_(G)/u_(L), wherein u_(G)and u_(L) are the superficial flow velocity of gas and liquid phases inm/s, respectively, and further wherein:Fr=[(L+G)υ_(LG)]² /gD _(h)We=D _(h)(L+G)²υ_(LG)/σ_(L)Re=D _(h)(L+G)/μ_(LG)υ_(LG)=(L/G)υ_(L)/(1+L/G)+υ_(G)/(1+L/G)μ_(LG)=(L/G)μ_(L)/(1+L/G)+μ_(G)/(1+L/G)D _(h)=2εD/[2+3(1−ε)( D/D _(p)) where D=column diameter, m D_(h)=bedhydraulic diameter, m D_(p)=equivalent particle diameter of catalyst, mFr=Froude number, unitless G=superficial mass velocity of gas, kg/m²sg=acceleration due to gravity, m/s² L=superficial mass velocity ofliquid, kg/m²s Re=Reynolds number, unitless u_(G)=ideal superficialvelocity of gas at reactor temperature and pressure, m/su_(L)□=superficial velocity of liquid, m/s We=Weber number, unitlessε=void fraction of catalyst bed, unitless μ_(G)=viscosity of gas, kg/m−sμ_(L)=viscosity of liquid, kg/m−s μ_(LG)=effective viscosity of thegas-liquid mixture, kg/m−s ρ_(G)=ideal gas density at reactortemperature and pressure, kg/m³ ρ_(L)=liquid density, kg/m³ σ_(L)=liquidsurface tension, N/m υ_(G)=ideal specific volume of gas at reactortemperature and pressure, m³/kg, υ_(L)=specific volume of liquid, m³/kgυ_(LG)=specific volume of the gas-liquid mixture, m³/kg
 8. The processaccording to claim 1, wherein said process further comprises passing afeed solution comprising an organic feedstock downwardly in co-currentwith a hydrogen-containing gas through a hydrogenation zone comprising abed of a hydrogenation catalyst, and wherein said organic feedstockcomprises at least one aldehyde or ketone selected from C3-C13 aldehydesand ketones.
 9. A process in a fixed bed catalytic reactor operatingunder the following conditions:Ta<N; and   (a)φ>a+(b·Ta);   (b) wherein variables N, a, and b are predetermined from agraph of plural (Ta, φ) coordinates according to claim
 1. 10. Theprocess according to claim 9, wherein N, a, and b are predetermined bydrawing a line separating the coordinates (Ta, φ) providing acceptableresults from those coordinates (Ta, φ) providing unacceptable results,said line having a y-intercept defined by the parameter “a” and a slopedefined by the parameter “b” in equation (b) and having a limit ofacceptable results defined by the parameter “N” in said equation. 11.The process according to claim 10, said process comprising passing afeed solution comprising an organic feedstock downwardly in cocurrentwith a hydrogen-containing gas through a hydrogenation zone comprising abed of a hydrogenation catalyst, under hydraulic conditions that satisfythe following equations:10<Ta<500; and   (a){0.045+(0.00035×Ta)}<φ<0.8.   (b)
 12. A process comprising passing afeed solution comprising an organic feedstock downwardly in cocurrentwith a hydrogen-containing gas through a hydrogenation zone comprising abed of a hydrogenation catalyst, monitoring said process and takingmeasurements so as to calculate whether or not the following equationshold in at least one point in the reactor:Ta<500; and   (a)φ>a+(b ·Ta);   (b) wherein variables a and b are predetermined, andwherein φ is the gas to liquid volumetric flow ratio, and Ta is theTalmor number defined by the following equation:Ta=(1+1/Fr)/(We+1/Re);   (c) wherein Fr is the Froude number, We is theWeber number, and Re is the Reynolds number.
 13. The process accordingto claim 12, further including calculating whether or not said equationshold for said reaction.
 14. The process according to claim 13, whereinif said equations do not hold for said reaction, changing at least onereactor variable selected from reactor diameter, reactor length,pressure, liquid rate, gas rate, gas/liquid stoichiometry, catalystparticle size, catalyst particle shape, catalyst loading density,catalyst bed void fraction, presence or absence of recycle, recyclestoichiometry, number of vessels, and heat removal method.
 15. Theprocess according to claim 14, further including, after said changing ofat least one reactor variable, repeating the steps of taking saidmeasurements and calculating whether or not said equations hold for saidreaction.
 16. An hydrogenation process in a downflowing fixed bedreactor comprising hydraulic conditions defined by the followingequations:10<Ta<500; and   (a){0.045+(0.00035×Ta)}<φ<0.8;   (b) wherein φ is the volumetric gas toliquid flow ratio and Ta is defined by the following equation:Ta=(1+1/Fr)/(We+1/Re);   (c) wherein Fr is the Froude number, We is theWeber number, and Re is the Reynolds number.
 17. A method of producingan alcohol comprising feeding a liquid comprising an organic speciescomprising at least 3 carbon atoms and a reactive carbonyl moiety into afixed bed reactor under conditions sufficient to cause hydrogenation ofsaid reactive carbonyl moiety, said conditions including hydraulicconditions determined according to claim
 1. 18. In a process forcarrying out a chemical reaction in a fixed bed reactor, the processcomprising providing at least one zone having a solid catalyst, a liquidphase, and a gaseous phase, under conditions sufficient to carry out achemical reaction, the improvement comprising operating said zone underhydraulic conditions obtainable according to the method of claim
 1. 19.The process according to claim 18, wherein said hydraulic conditions areobtained by obtaining new (Ta, φ) coordinates and comparing said newcoordinates to the boundary conditions obtained from said plot of (Ta,φ) coordinates, wherein said new coordinates are obtained by changing atleast one of hydraulic diameter, reactor diameter, reactor length,pressure, liquid rate, gas rate, gas/liquid stoichiometery, catalystparticle size, catalyst particle shape, catalyst loading density, numberof vessels, and heat removal method.
 20. The process according to claim19, wherein said process is repeated until (Ta, φ) coordinates areobtained wherein no free H₂ is observed at the end of the reactor. 21.The process according to claim 20, wherein heat removal methods areselected from adiabatic heat removal, heat removal by recycle withexternal cooling wherein the cooled recycle is injected in one or morepoints of the reactor, or by coils or tubes within the catalyst bed. 22.In a process for carrying out a chemical reaction in a fixed bed reactorcomprising determining the operating conditions required for saidchemical reaction, carrying out said chemical reaction in said fixed bedreactor, and obtaining at least one product from said chemical reaction,the improvement comprising: plotting plural (Ta, φ) coordinates on agraph, each coordinate obtained by carrying out said chemical reactionin said fixed bed reactor, wherein Ta is the inertia and gravity forcesdivided by the interface and viscous forces at a preselected point insaid reactor and φis the volumetric gas to liquid ratio at saidpreselected point, and obtaining from said graph boundary conditionsrequired for said operating conditions required for said chemicalreaction, and operating said chemical reaction within said boundaryconditions.
 23. A fixed bed catalytic reactor having at least one ofcolumn diameter, column length, bed hydraulic diameter, equivalentparticle diameter of catalyst, catalyst shape, and void fraction ofcatalyst bed determined by a process according to claim
 1. 24. A graphcomprising a plot prepared by the process of claim 1.